Apparatus and process for production of high purity hydrogen

ABSTRACT

The invention relates to a new and improved process and apparatus for the production of high purity hydrogen by steam reforming. The apparatus is an integrated flameless distributed combustion-membrane steam reforming (FDC-MSR) or reactor for steam reforming of a vaporizable hydrocarbon to produce H 2  and CO 2 , with minimal CO, and minimal CO in the H 2  stream. The flameless distributed combustion drives the steam reforming reaction which pro-vides great improvements in heat exchange efficiency and load following capabilities. The reactor may contain multiple flameless distributed combustion chambers and multiple hydrogen-selective, hydrogen-permeable, membrane tubes. The feed and reaction gases may flow through the reactor either radially or axially. A further embodiment of the invention involves producing high purity hydrogen by dehydrogenation using an integrated FDC-membrane de-hydrogenation reactor. A still further embodiment of the invention involves a zero emission hybrid power system wherein the produced hydrogen is used to power a high-pressure internally manifolded molten carbonate fuel cell. In addition, the design of the FDC-SMR powered fuel cell makes it possible to capture good concentrations of CO 2  for sequestration or use in other processes.

FIELD OF THE INVENTION

This invention relates to a process and apparatus for the production ofhigh purity hydrogen by steam reforming, to the separation of hydrogenproduced therefrom, and to the use of the hydrogen in a zero emissionhybrid power system incorporating a fuel cell.

BACKGROUND OF THE INVENTION

The production of electrical power in the most efficient manner withminimization of waste is the focus of much research. It would bedesirable to improve efficiency in the production of electricity,separate and use by-product CO₂ in other processes, and produce minimalNO_(x). The wide availability of natural gas with the highest H:C ratio(4:1) of any fossil fuel makes it a prime candidate for electricityproduction with minimum CO₂ emissions.

Electricity can be produced in fuel cells using pure hydrogen. Hydrogenproduction is commercially proven, but expensive. One method ofproducing hydrogen is steam methane reforming where hydrocarbons andwater are reacted to form CO and H₂, followed by a separatewater-gas-shift reaction where CO is reacted with H₂O to form CO₂ andH₂. The commercial application of these reactions in many refineriescommonly involves a series of reactors including a steam reformingreactor, and several post reactors to address the production of CO inthe reformer. The post reactors include a high temperature shiftreactor, a low temperature shift reactor, and a CO₂ absorber separator.Water and CO₂ separation is necessary to achieve pure hydrogen. Thereforming reactor is run at high pressure to avoid hydrogenrecompression downstream. The pressure lowers the equilibrium conversionsince reforming produces a positive net mole change. The steam reformingreaction is very endothermic; and the shift reaction is also exothermic.The conventional steam reforming reactors are operated above 900° C. topush the equilibrium toward complete formation of CO and H₂. The hightemperature causes severe corrosion and stress problems on theequipment. Steam reforming reactors are generally large to accomplisheconomies of scale. In addition, the typical operation of the shiftreactor at a lower temperature than the reforming reactor makes itimpractical to combine these two chemical reactions in one reactor.Furthermore, designs currently known do not lend themselves to beingscaled down to a smaller size or to making it possible to efficientlycontrol the temperature at various points.

Even if a reactor was capable of producing only CO₂ and H₂ and theconventional post reactors could be eliminated, the issue of CO₂separation would remain.

It would be desirable in the art to provide a steam reformer reactordesign for producing hydrogen substantially free of carbon and carbonoxides and with minimal production of NO_(x). If the high purityhydrogen produced could be used to create power in a hybrid system thatcould be compact in design and provide greater efficiency, such as morethan 71%, in the production of energy it would represent a distinctadvance in the art. In addition, it would be desirable if lowertemperatures could be used and if the entire process permitted morecontrol over temperatures at various points, or load-followingcapabilities. It would also be desirable to provide the modularityneeded at bulk-hydrogen production scales so that a producer can matchthe desired capacity by installing multiple reactor units of thespecific design. This is more cost-effective than either trying to scaleup or down the existing large box furnace reactor designs or buildingseveral thousand single-tube reactors. It would also be desirable toemploy less volume than conventional processes by intensifying theprocess and using less catalyst and smaller heater space. Furthermore,if the process produced CO₂ in higher concentrations and greater puritythan other processes in the art, and the CO₂ could be sequestered forother uses, it would be extremely desirable. Such an integrated systemwould demonstrate far greater efficiency than any power generatingsystem currently available.

SUMMARY OF THE INVENTION

The invention relates to an improved process and apparatus for theproduction of high purity hydrogen by steam reforming. The apparatus isan integrated flameless distributed combustion-membrane steam reforming(FDC-MSR) reactor for steam reforming of a vaporizable hydrocarbon toproduce H₂ and CO₂, with minimal CO as end product, and minimalconcentration of CO in the H₂ stream. The reactor may contain multipleflameless distributed combustion chambers and multiplehydrogen-selective, hydrogen-permeable, membrane tubes. The feed andreaction gases may flow through the reactor either radially or axially.A further embodiment of the invention involves an integrated flamelessdistributed combustion membrane dehydrogenation reactor fordehydrogenation of a hydrocarbon-containing compound, such as ethylbenzene, to form hydrogen. A still further embodiment of the inventioninvolves a zero emission hybrid power system wherein the producedhydrogen is used to power a high-pressure internally manifolded moltencarbonate fuel cell. In addition, the design of the FDC-SMR powered fuelcell makes it possible to capture good concentrations of CO₂ forsequestration or use in other processes such as for enhanced oilrecovery.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic diagram of the novel membrane steam-reforming(MSR) reactor with a flameless distributed combustion (FDC) heatersection, catalyst section, and permeate section placed in order from theoutside in.

FIG. 2 is a schematic diagram of another embodiment of the novel FDC-MRSreactor of the invention.

FIG. 3 is a graph showing molar fraction and methane conversion alongthe reactor.

FIG. 4 is a graph showing temperature and heat flux per length profilealong the reactor.

FIG. 5 is a graph showing hydrogen molar fraction profile and membranevolumetric flux per length (in m³/m/s) along the reactor.

FIG. 6 is a simplified flow diagram of the zero emission flamelessdistributed combustion membrane steam reformer fuel hybrid power system.

FIGS. 7A&B is a process flow diagram of the zero emission process assimulated in a HYSYS process simulator.

FIG. 8 is a schematic diagram of a multi-tubular, FDC heated, radialflow, membrane, steam reforming reactor in accordance with theinvention. Some of the inlet and outlet streams of the membrane and FDCtubes have been omitted for simplicity.

FIG. 9 is a cross section of the shell of the multi-tubular, FDC heated,radial flow, membrane reactor shown in FIG. 8.

FIGS. 10A and 10B are schematic diagrams of a “closed ended” and of an“open ended” FDC tubular chamber used to drive the reforming reactionsin the process and apparatus of the present invention.

FIG. 11 is a schematic diagram of a multi-tubular, FDC heated, axialflow, membrane steam reforming reactor in accordance with the invention.

FIGS. 12 is a cross section of the shell of the Multi-tubular, FDCheated, axial flow, membrane reactor shown in FIG. 11.

FIGS. 13A & 13B and 13C & 13D are schematic diagrams of two baffleconfigurations which can be employed to increase the contact of thereactant gases with the catalyst in a multi-tubular, FDC-heated, axialflow, membrane reactor in accordance with the invention.

FIGS. 14, 15 16 and 17 are top cross section views of the shells ofother embodiments of the multi-tubular, FDC heated, axial flow,membrane, steam reforming reactors of the invention.

DETAILED DESCRIPTION OF THE INVENTION

The present invention provides a new process and apparatus for steamreforming of any vaporizable hydrocarbon to produce H₂ and CO₂, withminimal CO as endproduct, and having minimum concentration of CO in theH₂ stream, said process being accomplished in one reactor, at lowertemperatures than those used in conventional steam methane reformingreactors, constantly removing pure hydrogen, and using as a heat sourceflameless distributed combustion which provides great improvements inheat exchange efficiency and load-following capabilities to drive thesteam reforming reaction. Similar efficiency and load-following issimply not possible with conventional firebox steam reformer furnacedesigns and multi-reactor shift units. The flameless distributedcombustion heat source makes it possible to transfer between 90 and 95%of the heat to the reacting fluids. In another embodiment, the inventionis also a zero emission hybrid power system wherein the producedhydrogen is used to power a high-drogen is used to power a high-pressureinternally or externally manifolded molten carbonate fuel cell. Thesystem is capable of achieving 71% or greater efficiency in theconversion of fuel to electricity. In addition, the design of thisflameless distributed combustion—membrane steam reforming reactor(FDC-MSR) fueled hybrid system makes it possible to capture highconcentrations of CO₂ for sequestration or use in other processes.Finally, the design of the system can be scaled down to a mobile,lightweight unit.

Moreover, at bulk-hydrogen production scales, the multi-tubular(multiple FDC tubes and multiple hydrogen selective and permeablemembrane tubes) containing reactor disclosed herein provides themodularity needed. A producer can match the desired capacity byinstalling multiple reactor units of the specific design or havingmultiple FDC tubes and/or multiple hydrogen selective and permeablemembrane units in a large steam reformer. This is more cost-effectivethan either trying to scale up or down the existing large box furnacereactor designs or building several thousand single-tube reactors.

The process for steam reforming of any vaporizable hydrocarbon toproduce purified H₂ and CO₂ comprises:

-   -   a) Providing a generally tubular reforming chamber having one or        more inlets for vaporizable hydrocarbons and steam, and one or        more corresponding outlets for byproduct gases, including H₂O        and CO₂, with a flow path in between said inlet and outlet, and    -   having one or more inlets for sweep gas (which may be H₂O in the        form of steam, or other gas such as recycled CO₂, nitrogen or        condensable hydrocarbons) and corresponding outlets for the        sweep gas and hydrogen, with a flow path between said inlet and        outlet, and    -   having one or more inlets for preheated air and corresponding        inlets for fuel gas mixtures, with a flow path between said        inlets containing at least one, particularly plurality of        flameless distributed combustion heaters,

wherein said flow path for vaporizable hydrocarbon and flow path forsweep gas form two concentric sections with an annulus between having areforming catalyst therein;

b) Feeding a vaporizable hydrocarbon and steam into said reformingchamber through said inlet for a vaporizable hydrocarbon and steam;

c) Flowing said vaporizable hydrocarbon over a reforming catalyst;

d) Causing both steam reforming-and the shift reaction to take place insaid reforming chamber; and

e) Conducting said reforming in the vicinity of at least onehydrogen-permeable and hydrogen-selective membrane, whereby purehydrogen permeates said membrane;

f) Wherein heat to drive said reaction is provided by said flamelessdistributed combustors.

The process of the present invention may also be described as a processfor the production of hydrogen, comprising:

-   -   a) reacting steam with a vaporizable hydrocarbon at a        temperature of from about 200° C. to about 700° C. and at a        pressure of from about 1 bar to about 200 bar in a reaction zone        containing reforming catalyst to produce a mixture of primarily        hydrogen and carbon dioxide, with a lesser amount of carbon        monoxide;    -   b) providing heat to said reaction zone by employing one or more        flameless distributed combustion chambers (which may in the form        of a tube or other shape) thereby driving said reaction;    -   c) conducting said reaction in the vicinity of one or more        hydrogen-permeable and hydrogen-selective membranes, which may        be in the form of a tube or others, whereby hydrogen formed in        said reaction zone permeates through said selective membrane(s)        and is separated from said carbon dioxide and carbon monoxide.

In order to produce electricity with zero emissions and capture CO₂, thehigh purity hydrogen which permeates the membrane may be directed to theanode of a high pressure molten carbonate fuel cell and the by-productsfrom the reforming reaction are directed to the cathode of said fuelcell. The high purity hydrogen can also be directed to other types offuel cells, such as PEM (proton exchange membrane) fuel cells or SOFC(solid oxide fuel cells) and the like.

The invention also pertains to an apparatus comprising a membrane steamreformer heated by flameless distributed combustion to produce hydrogenthat may be used for a variety of purposes including as fuel to a fuelcell such as high pressure molten carbonate fuel cell or a PEM fuelcell. The integrated flameless distributed combustion-membrane steamreforming reactor of the present invention comprises:

-   -   A reforming chamber comprising a generally tubular reactor        having two concentric sections comprising a larger outside        section and a smaller inside section and an annulus between said        sections, wherein said outside section has an inlet for        preheated air and a corresponding inlet for fuel gas, with a        flow path between and a plurality (two or more) of flameless        distributed combustors arranged in a circular path in said        outside section; and wherein said inside section has an inlet        for sweep gas and an outlet on said opposite end for sweep gas        and H₂, and said annulus has an inlet for vaporizable        hydrocarbons and an outlet for by-product compounds and a        perm-selective (hydrogen selective), hydrogen-permeable membrane        positioned either on the inside or outside of the annular        section.

The present invention also pertains to a flameless distributedcombustion (FDC) heated, membrane, steam reforming reactor comprising:

a) a reforming chamber containing a reforming catalyst bed, thereforming chamber having an inlet for vaporizable hydrocarbon and steam,a flow path for hydrogen and by-product gases resulting from thereforming reactions taking place in the reforming chamber and an outletfor the by-product gases,

b) at least one flameless distributed combustion (FDC) chamber in a heattransferring relationship with the reforming catalyst bed whereby adistributed, controlled heat flux is provided by the FDC chamber to thereforming catalyst bed, said FDC chamber comprising an inlet and a flowpath for an oxidant, an outlet for combustion gas and further comprisinga fuel conduit having an inlet for fuel and a plurality of fuel nozzleswhich provide fluid communication from within the fuel conduit to theflow path of the oxidant, the plurality of fuel nozzles being sized andspaced along the length of the fuel conduit so that no flame resultswhen the fuel is mixed with the oxidant in the FDC chamber;

c) a preheater for preheating the oxidant and/or fuel to a temperaturethat when the fuel and oxidant are mixed in the FDC chamber, thetemperature of the resulting mixture of oxidant and fuel will exceed theautoignition temperature of said mixture; and

d) at least one hydrogen-selective, hydrogen-permeable, membrane tube incontact with the reforming catalyst bed, the membrane tube having anoutlet whereby hydrogen formed in the reforming chamber permeates intosaid membrane tube and passes

through said outlet.

The present invention also relates to a flameless distributed combustionheated, membrane, dehydrogenation reactor comprising:

a) a dehydrogenation chamber containing a catalyst bed, saiddehydrogenation chamber having an inlet for vaporizable hydrocarbon, aflow path for hydrogen and product gases resulting from thedehydrogenation reactions taking place in said dehydrogenation chamberand an outlet for said product gases,

b) at least one flameless distributed combustion chamber in a heattransferring relationship with said catalyst bed whereby a distributed,controlled heat flux is provided by said flameless distributedcombustion chamber to said catalyst bed, said flameless distributedcombustion chamber comprising an inlet and a flow path for an oxidant,an outlet for combustion gas and further comprising a fuel conduithaving an inlet for fuel and a plurality of fuel nozzles which providefluid communication from within the fuel conduit to the flow path ofsaid oxidant, said plurality of fuel nozzles being sized and spacedalong the length of said fuel conduit so that no flame results when saidfuel is mixed with said oxidant in said flameless distributed combustionchamber;

c) a preheater capable of preheating said oxidant to a temperature thatwhen said fuel and said oxidant are mixed in said flameless distributedcombustion chamber, the temperature of the resulting mixture of saidoxidant and fuel exceeds the autoignition temperature of said mixture;and

d) at least one hydrogen-selective, hydrogen-permeable, membrane tube incontact with said catalyst bed, said membrane tube having an outletwhereby hydrogen formed in the dehydrogenation chamber permeates intosaid membrane tube and passes through said outlet.

The present invention further relates to process for dehydrogenation ofethylbenzene, which process comprises the steps of feeding ethylbenzeneinto the reactor as described above to produce styrene and hydrogen. Thecatalyst bed contains a dehydrogenation catalyst such as an ironoxide-containing catalyst.

In a preferred embodiment of the invention, the aforesaid FDC heated,membrane, steam reforming reactor contains multiple FDC chambers(preferably, but not necessarily, in the form of tubes) and multiplehydrogen-selective, hydrogen-permeable membrane tubes disposed in, orotherwise in contact with, the reforming catalyst bed in the reformingchamber. Examples of multi-tubular reactors in accordance with theinvention are shown in FIGS. 8-9, 11-12, and 14-17.

The multi-tubular, FDC heated, membrane, steam reforming reactors inaccordance with the invention may be either of the radial flow type asshown in FIGS. 8 and 9, or may be of the axial flow type as shown inFIGS. 11-12 and 14-17. In a radial flow reactor the gases generally flowthrough the reforming catalyst bed radially from outside to inside (orfrom inside to outside), while in an axial flow reactor the gasesgenerally flow through the reforming catalyst bed in the same directionas the axis of the reactor. In the case of a vertical reactor, the flowwould be from the top of the reactor to the bottom, or the bottom of thereactor to the top.

The multi-tubular, FDC heated, membrane steam reforming reactors inaccordance with the present invention may contain from as few as 2 FDCtubes up to 100-or more, particularly 3 to 19, depending the size of theFDC tubes, the size of the catalyst bed and the level of heat fluxdesired in the catalyst bed. The size of the FDC tube can vary fromabout 1 inch OD up to about 40 inches or ore OD. The number ofhydrogen-selective membrane tubes may also vary from as few as 2 up to400 or more, particularly 3 to 90. The size of the membrane tubes mayvary from about 1 inch up to about 10 inches or more. In general, theratio of FDC tube surface area to membrane tube surface area will be inthe range of about 0.1 to about 20.0, particularly from about 0.2 toabout 5.0, more particularly from about 0.5 to about 5.0, still moreparticularly from about 0.3 to about 3.0 and even more particularly fromabout 1.0 to about 3.0. The term “surface area” when used in referenceto the above ratios, means the external (circumferential) area of theFDC tubes and the membrane tubes. For instance, a 1 inch OD tube of 12inches length would have an external surface area of 37.6 square inches.

Each FDC tube or chamber will have at least one fuel conduit disposedtherein. Larger FDC chambers generally will have multiple fuel conduits.The FDC chambers or tubes employed in the multi-tubular reactors of theinvention may be “open ended” or “closed ended” as discussed below inconnection with FIGS. 10A and 10B.

A sweep gas may be used to promote the diffusion of hydrogen through thehydrogen-selective, hydrogen-permeable membrane. In case a sweep gas isemployed, the membrane tube may contain an inlet and flow path for sweepgas feed and a flow path and outlet for the return of sweep gas andpermeated hydrogen.

Baffles and/or screens may also be employed in the multi-tubularreactors of the present invention to improve contact of the reactivegases with the catalyst and to improve flow distribution. The FDC tubesand/or membrane tubes may also be surrounded by cylindrical screens toprotect the tubes from direct contact with the catalyst.

In a further embodiment of the invention the reforming chamber of areactor in accordance with the invention is in communication with a highpressure molten carbonate fuel cell, wherein the outlet for hydrogenfrom the reformer is in communication with the anode of said fuel celland the outlet for by-product compounds is in communication with thecathode of said fuel cell.

The integrated FDC-MSR process and apparatus of this invention iscapable of producing high purity hydrogen with minimal production of CO,particularly less than about 5 molar %, more particularly less than 3molar %, and still more particularly less than 2 molar % on a molar drybasis of the total products, and with less than 1000 ppm of CO andparticularly less than 10 ppm of CO on a dry basis, more particularlyvirtually no CO in the hydrogen stream produced. By practice of thepresent invention it is possible to produce high purity hydrogen e.g.,hydrogen having a purity on a dry basis of greater than 95%. The presentinvention can be used to produce hydrogen having purities as high as97%, 99%, or under optimum conditions 99+%. The effluent (by product)stream from the MSR reactor will typically contain more than 80% CO₂ ona dry basis, e.g., 90% CO₂, 95% CO₂ or 99% CO₂, and less than about 10%CO on a dry basis, e.g., less than about 5% CO, preferably less than 1%Co.

Total heat management and turbines may be included in the system toincrease the efficiency and produce additional electricity or to douseful work such as compress gases or vapors.

One aspect of the present invention is a flameless distributedcombustion heated membrane steam reformer hydrogen generator. In thedesign of the invention there are disclosed distinct improvements inoverall efficiency, particularly size, scalability and heat exchange.The present invention typically employs only one reactor to produce thehydrogen versus typically four reactors used in conventional processes,and part of the heat load is supplied by the water-gas-shift reaction.The design of the invention captures essentially all of the heat in thereaction chamber since heat exchange occurs on a molecular level, whichreduces the overall energy requirements.

Chemical equilibrium and heat transfer limitations are the two factorsthat govern the production of hydrogen from methane in conventionalreactors. These factors lead to the construction of large reactorsfabricated from expensive high temperature tolerant materials. They areenclosed in high temperature furnaces that are needed to supply the highheat fluxes.

In the present invention the two major limitations of chemicalequilibrium and heat transfer are overcome by the innovative combinationof an in-situ membrane separation of hydrogen in combination with aflameless heat source comprising flameless distributed combustion (FDC)that makes it possible to more efficiently use all the energy in thesystem, as well as provide load following capabilities.

The reformer of the present invention reduces the operating temperatureof the steam reforming reactor close to the lower temperature used in ashift reactor. With the temperatures for the steam reforming and shiftcloser, both operations are combined into one reactor. With bothreactions occurring in the same reactor the exothermic heat of reactionof the shift reaction is completely captured to drive the endothermicsteam reforming reaction. This reduces the total energy input for thesum of the reactions by 20%. The lower temperature reduces stress andcorrosion and allows the reactor to be constructed from much lessexpensive materials. Combining the operations also reduces the capitaland operating cost since only one reactor, instead of two or three, arerequired. Moreover, the reaction is not kinetics-limited even at thelower temperature, thus, the same or even less catalyst can be used.

The general description for steam reformers, including but not limitedto the reactions, enthalpies, values of equilibrium constants,advantages of integrated FDC-SMR reactor, as well as the advantages ofthe use of the membrane in the reactor can be found in US 2003/0068269,the entire descriptions of which are herein incorporated by reference.

The in-situ membrane separation of hydrogen employs a membranefabricated preferably from an appropriate metal or metal alloy on aporous ceramic or porous metal support, as described below, to drive theequilibrium to high conversions. With constant removal of the hydrogenthrough the membrane, the reactor can be run at much lower than thecommercially practiced temperatures of 700-900+° C. A temperature of500° C. is sufficient to drive the kinetics to high conversions when theequilibrium is shifted using the hydrogen separation membrane. At thistemperature the selectivity to CO₂ is almost 100%, while highertemperatures favor the formation of CO as a major product.

FIG. 1 shows a schematic diagram of a membrane steam-reforming reactorwith a flameless distributed combustion (FDC) heater section, catalystsection, and permeate section. The reactor 1 shown in FIG. 1 consists oftwo concentric sections. The outer concentric section 2 is the FDCheater section, while the inner concentric section 3 is the permeatesection. The annulus, 4, in between is the catalyst section. The term“reforming catalyst” as used herein means any catalyst suitable forcatalyzing a steam reforming reaction, which includes any steamreforming catalyst known to one skilled in the art, as well as any“pre-reforming catalyst” which is suitable for catalyzing steamreforming reactions in addition to being suitable for processing heavierhydrocarbons prior to a steam reforming reaction. Reforming catalyst isloaded into the annulus section 4 wherein the above-described reactionstake place. (section 4 is also variously referred as the catalystsection, the reaction section or the reaction zone). The membrane, 8, isrepresented on the inside of the small section, 3, (the permeatesection) in FIG. 1. The FDC fuel tubes, 10, are placed in a circularpattern in the FDC heating section, while the air flows in that annularregion surrounding the fuel tubes. While FIG. 1 shows the FDC heatersection, catalyst section and permeate section placed in order fromoutside in, the location of the membrane and FDC heater section can bereversed to achieve higher membrane area.

The feed gas stream containing a mixture of vaporizable hydrocarbon(e.g. naphtha, methane or methanol) and H₂O with a minimum overall O: Cratio of 2:1 enters catalyst section 4 at 5. If used, sweep gas forpromoting the diffusion of hydrogen through the membrane enters the topof the permeate section 3 at 6. Alternatively, sweep gas can beintroduced into the permeate section by means of a stinger pipe fittedto bottom of the permeate section. In case of this alternative, hydrogenin sweep gas would exit the permeate zone at the bottom of the permeatesection at. 12. Optionally, the stinger pipe to introduce the sweep gasmay be connected at the top of the permeate section in which case thehydrogen and sweep gas would exit at the top of this section. Preheatedair enters the FDC heater section at 7. Hydrogen (pure or in sweep gas)exits at 12. Flue gas from the FDC heater section exits at 11. Unreactedproducts and by-products (e.g., CO₂, H₂O, H₂, CH₄, and CO) exit catalystsection 4 at 13. Fuel 14 (which may include a portion of the hydrogenexiting the permeate section or part of the reactor effluent) enters theFDC fuel tubes 10 as shown and is mixed with the preheated air in theFDC heating section. It is also possible to remove the produced hydrogenusing a vacuum instead of a sweep gas.

FIG. 2 shows a schematic diagram of another embodiment of the integratedFDC-MSR reactor of the present invention. The reactor depicted in FIG.2, similar to the reactor in FIG. 1, has an outer concentric FDC heatersection 2 and an inner permeate section 3, with an intermediate catalystsection 4 containing catalyst 9. The catalyst section also contains alayer of inerts 15 at the top of the catalyst section. A feed streamcontaining a vaporizable hydrocarbon (e.g. naphtha, methane or methanol)and steam enter the reactor at 5, while sweep gas, if used, enters thereactor at 6. Fuel for the FDC heater section enters fuel tubes 10 at14. However, in the case of this embodiment the fuel enters the FDCheating section at the top of the section and flows concurrently withthe preheated air (or other oxidant) which enters the FDC heatingsection at 7. The flow in the FDC heating section is also co-currentwith the flow of the reactant gases in section 4. Fuel tubes 10 have aplurality of openings or nozzles sized and spaced along the length ofthe fuel tubes so that the amount of fuel mixing with the air or oxidantin the annular part of the FDC section surrounding the fuel tubes can becontrolled to achieve the desired heat distribution along the length ofthe FDC heating section which in this embodiment surrounds the reactionsection. Flue gas containing very low levels of NO_(x) leaves the FDCheater section at 11, while effluent from the catalyst (reaction)section exits at 13. Hydrogen formed in reaction section permeatesthrough hydrogen-selective, hydrogen-permeable, membrane 8 and exits thepermeate section (as such or with sweep gas) at 12.

The novel integrated FDC-membrane steam-reforming reactor of the presentinvention operates at a lower temperature than that used in conventionalsteam methane reformers. A suitable temperature is less than about 700°C., for example in the range of from about 300 to about 650° C. In somecases lower temperatures, e.g., as low as about 200° C. can be used upto about 600° C. A preferred range is from about 400 to about 550° C.,more preferably from about 400° C. to about 500° C. Suitable pressure isin the range of from about 1 to about 200 bar, preferably from about 10to about 50 bar. The simulation in Example 1 of the present inventionwas carried out at a temperature of about 500° C. and 30 Bar. This lowtemperature achieves high selectivity to CO₂ and negligible selectivityto CO.

Any vaporizable (or optionally oxygenated) hydrocarbon can be used inthe present process and apparatus, including, but not limited to,methane, methanol, ethane, ethanol, propane, butane, light hydrocarbonshaving 1-4 carbon atoms in each molecule, and light petroleum fractionslike naphtha at boiling point range of 120-400° F., which is a typicalfeed for commercial steam reformers. Petroleum fractions heavier thannaphtha can also be employed like diesel or kerosene or jet fuel atboiling point range of 350-500° F. or gas oil at boiling point range of450-800° F. Hydrogen, carbon monoxide and mixtures thereof, e.g.,syngas, may also be used in the process and apparatus of the presentinvention, and are included in the definition of “vaporizablehydrocarbon”. Methane was used in the examples to demonstrate theprocess.

With the FDC-MSR process and apparatus of the present invention it ispossible to use O: C ratios as low as 2.8, down to 2.6, without cokingproblems, with the minimum O: C ratio being about 2:1. This resultslower energy costs if methane is used as the feed in the presentinvention, since lower steam to methane ratios can be used thusrequiring less energy to vaporize water. Because of the ability tooperate at lower O:C ratios, it is also possible to use heavier, lessexpensive feeds in the FDC-MSR reactor of the present invention than canbe used in conventional steam methane reformers.

In another embodiment of the invention, the integrated FDC-MSR processand apparatus of the invention can be used to perform water-gas-shiftreactions on syngas mixtures (i.e., mixtures of hydrogen and carbonmonoxide) produced from conventional processes like Catalytic PartialOxidation (CPO), Steam Methane Reforming (SMR) and Autothermal Reforming(ATR). The integrated FDC-MSR reactor is well suited for this since itproduces high purity hydrogen and converts carbon monoxide to carbondioxide and more hydrogen. Thus, the versatile FDC-MSR reactor of theinvention is capable of replacing the high temperature shift, lowtemperature shift and methanation reactors and the hydrogen purificationsection. A mixture of syngas and vaporizable hydrocarbon can also beused to yield a net reaction which may be either endothermic, thermallyneutral or slightly exothermic.

The reactor annulus is packed with steam reforming catalyst and equippedwith a perm-selective (i.e., hydrogen selective) membrane that separateshydrogen from the remaining gases as they pass through the catalyst bed.The steam reforming catalyst can be any known in the art. Typicallysteam reforming catalysts which can be used include, but are not limitedto, Group VIII transition metals, particularly nickel. It is oftendesirable to support the reforming catalysts on a refractory substrate(or support). The support is preferably an inert compound. Suitablecompounds contain elements of Group III and IV of the Periodic Table,such as, for example the oxides or carbides of Al, Si, Ti, Mg, Ce andZr. The preferred support composition for the reforming catalyst isalumina.

The catalyst used in the examples to demonstrate the present inventionwas nickel on porous alumina. As the hydrogen is formed in the catalystbed, it is transported out through the hydrogen-permeable separationmembrane filter. Advantages of this technology include the capacity toseparate essentially pure hydrogen from any poisons that may also bepresent, including CO and H₂S, and from other fuel diluents. The poisonsdo not pass through the separation membrane, which is fabricated fromone of a variety of hydrogen-permeable and hydrogen-selective materialsincluding ceramics, carbon, and metals.

Membranes which are suitable for use in the apparatus and process of thepresent invention include, but are not limited to, (i) various metals,such as hydrogen permeable transition metals selected from Group IIIB,IVB, VB, VIIB and VIIIB of the periodic table and metal alloys or metalhydrides of such metals, (ii) molecular sieves, ceramics, zeolites,silica, alumina, refractory metal oxides, carbon, (iii) organicpolymers, and mixtures thereof. Illustrative, but non-limiting, examplesof hydrogen separating devices utilizing such membranes include themembranes described in U.S. Pat. No. 5,217,506, issued Jun. 8, 1993 toDavid J. Edlund et al, U.S. Pat. No. 5,259,870, issued Nov. 9, 1993 toDavid J. Edlund et al and U.S. Pat. No. 5,451,386, issued Sep. 19, 1995to Collins et al, and U.S. Pat. No. 6,152,987, issued Nov. 28, 2000, thedescriptions of all of which are incorporated herein by reference.

Membranes, which are especially suitable for use in the presentinvention, include various metals and metal alloys on porous ceramic orporous metal supports. The porous ceramic or porous metallic supportprotects the membrane surface from contaminants and, in the formerchoice, from temperature excursions. Illustrative, but non-limitingexamples of materials suitable for use as a support for the membraneswhich may be used in the apparatus and process of the present inventioninclude an inorganic porous material such as palladium, platinum,palladium alloys, porous stainless steel, porous silver, porous copper,porous nickel, porous Ni-based alloys, metal mesh, sintered metalpowder, refractory metals, metal oxides, ceramics, porous refractorysolids, honeycomb alumina, aluminate, silica, porous plates, zirconia,cordierite, mullite, magnesia, silica matrix, silica alumina, porousVycar, carbon, glasses, and the like.

A particularly suitable membrane support is porous stainless steel orporous Ni-based alloy. Porous nickel-based alloys, like Hastelloy andInconel, are particularly suitable as being stable at high temperatures.Ni-based alloys have also high mechanical strength and this strength ismaintained at high temperatures. Ni-based alloys also have highresistance to oxidation and scaling when exposed to steam, a feed thatis present in steam reforming reactions. Ni-based alloys also have highresistance to chloride pitting. This assures that the support will notpit if there is trace of chloride left over from the plating solutionafter the rinsing and drying steps usually employed. Particularly, Alloy625 (or Inconel 625) is superior in resistance to crevice corrosion,uniform corrosion and stress corrosion cracking. It has a niobiumaddition that stabilizes the alloy against sensitization during welding,thereby preventing subsequent inter-granular attack. Alloy 625 isresistant to hydrochloric acid, nitric acid, neutral salts and alkalimedia. Alloy 625 resists cracking in both oxidizing and non-oxidizingenvironments. It has very high allowable design strength and is able towithstand temperatures up to 760° C. Alloy 625 has a high amount ofChromia (ceramic) on the surface, which can act as a barrier to theinter-metallic diffusion of Palladium with Fe (iron) or other metals.The Pd layer can be deposited on the outside of the porous ceramic ormetallic support, in contact with the catalyst section, or it can bedeposited on the inside thereof. The inertness, range of porosityavailable, and the fact that, to some extent, alumina can function as aninsulator, also make it a good choice for the support. Additionaladvantages include the fact the alumina can function to filter offmaterial that might deposit on the membrane and plug it. The use ofalumina also makes it possible to control the distance of the membranefrom the catalyst section, and, therefore, control the temperature dropacross the operating membrane at a given temperature and maximumeffectiveness, and lessens the likelihood of overheating. It is alsopossible to use the ceramic support as an insulating layer to keep themembrane at the design temperature. The temperature of the sweep gas mayalso be controlled to adjust the membrane temperature. The membranepermeate side can provide extra heat transfer area, with superheatedsteam used as sweep gas and also as a heat transfer fluid for heatingand temperature control. Also, a combustion catalyst with some oxygeninjection via a perforated tube can oxidize some of the producedhydrogen to supply the enthalpy to drive the steam reforming reaction.The presence of Pd or Pd-alloy in the vicinity of the air and hydrogenmixture makes this reaction occur at lower temperature than theautoignition temperature of hydrogen and air (which is 571° C.). Thisresults in a heat source that does not exceed the maximum operatingtemperature of the preferred Pd membrane, which is around 550° C. Thisinternal heating concept is based on flameless distributed combustionconcepts and is an example of inverse combustion and may be used with orwith out catalyst surrounding the perforated tube that supplies theoxygen. Optionally, a suitable methanation catalyst may be placed in thepermeate compartment as an extra safeguard to CO penetration through thehydrogen membrane if small pinholes develop. This catalyst could convertCO to methane and keep the CO level in the hydrogen stream always in theparts per million range. Typically, the CO level in the hydrogen streamexiting the membrane steam reforming reactor of the present inventionwill be less than about 10 ppm, e.g., less than 5 ppm, 2 ppm, 1 ppm or0.1 ppm.

Preferred materials for fabricating said membrane include mainly, butnot exclusively, metals of Group VIII, including, but not limited to Pd,Pt, Ni, Ag, Cu, Ta, V, Y, Nb, Ce, In, Ho, La, Au, etc. particularly inthe form of alloys. Pd and Pd alloys are preferred. The membrane used todemonstrate the present invention was a very thin film of a palladiumalloy having a high surface area. Membranes of this type can be preparedusing the methods disclosed in U.S. Pat. No. 6,152,987, which isincorporated by reference herein in its entirety. Platinum or a platinumalloy would also be suitable.

As mentioned, with respect to FIG. 1, the membrane is pictured on theinside of the smaller (i.e., the inner) concentric section, whichreduces the surface area to a minimum. In order to obtain greater flux,the membrane could be placed on the outside of the larger section of thereactor. Changes in geometry of the membrane permit a number of optionsdepending on requirements as will be apparent to those skilled in theart. For example, one option is to place the membrane on the outside ofthe reactor wall to achieve higher surface area. If in Example 1 themembrane was placed on the outside tube of the 14 cm diameter, thesurface area value can be increased by a factor of 2. Also, more tubesof smaller diameter can be used to achieve a higher surface to volumeratio. A jagged cross-section of the membrane tube (with star shape forinstance) could increase the surface area. Finally, the space velocityof the gas may be reduced, e.g. by 2-3 or 2200-3300 h⁻¹, to allow moretime for the hydrogen to diffuse through the membrane.

The hydrogen separation membrane used in Example 1 was a Palladium-alloy(such as alloy of palladium with one or more other metals such as Ag,Cu, Au, Ta, V, etc.) thin film of 1 μm or less with a high surface area.The Pd-alloy film is supported on a porous ceramic matrix that acts asthe mechanical support and a filtration medium to prevent coke fromcovering the membrane. The porous ceramic support also acts as aninsulator to reduce heat losses from the reactor. It also keeps themembrane at the specified temperature for optimum performance andstability. This special design geometry is highly efficient. Thepermeability used for the base case is 7.8 10⁻² std-m³/m²/s/Bar^(0.5)which is a number 2-30 times higher than reported in the literaturewhich can be found in Table 2 of US2003/0068269, the description thereofand the commercially available membranes described therein are hereinincorporated Steam is not known to cause a problem in membranestability, however, if any problems developed at higher temperatures,the replacement of water with recycled carbon dioxide or nitrogen assweep gas is a viable alternative. Other sweep gases could be used, likehydrocarbons, or mixtures thereof, with a moderate boiling point of100-400° C. These would condense at temperatures closer to the permeateoutlet temperature and thus reduce the energy loss during cooling andreheating of the sweep gas. Hydrocarbons have lower condensationenthalpy than water, thus, they may reduce the heat exchanger sizerequirements. They can also reduce the sweep gas impurities in thepurified H₂ stream since they have low vapor pressure at thecondensation temperature. A mixture of hydrocarbons can make thecondensation occur in a range of temperatures, and thus, avoid the pinchpoint limitation occurring with a single and sharp boiling point.

As a particular embodiment of the present invention, the permeatesection can be connected to a metal hydride precursor compartment whichreacts with the permeating hydrogen to form metal hydride. This reactionreduces the effective partial pressure of hydrogen in the permeatestream and increases the driving force for hydrogen flux.

In the present invention heat transfer limitations are overcome by theinnovative use of flameless distributed combustion (FDC) as the primaryheat source. FDC is used to distribute heat throughout the reactor athigh heat fluxes without high temperature flames and with low NO_(x)production. This is achieved by injecting small quantities of fuel intoa preheated air stream and reaching autoignition conditions. Fuelquantity is controlled by nozzle size, the temperature rise is verysmall, and there is no flame associated with the combustion (combustionis kinetically limited, rather than mass-transfer limited). The reactionin the case where methane is used as fuel for FDC is:

Combustion: CH₄+2O₂

CO₂+2H₂O −802.7 kJ/gmol

Comparing the enthalpies of this reaction with the reforming of methaneto CO₂, it is obvious that the minimum amount of methane that needs tobe combusted in order to support reforming is 17% of the total methaneused (ratio of 1:4.9 to the reformed methane).

Flameless distributed combustion is disclosed in U.S. Pat. No.5,255,742, U.S. Pat. No. 5,862,858, U.S. Pat. No. 5,899,269, U.S. Pat.No. 6,019,172, and EP 1 021 682 B1 the disclosures of which are herebyincorporated by reference herein in their entirety.

An important feature of the flameless distributed combustion is thatheat is removed along the length of the combustion chamber so that atemperature is maintained that is significantly below what an adiabaticcombustion temperature would be. This almost eliminates formation ofNO_(x), and also significantly reduces metallurgical requirements, thuspermitting the use of less expensive materials in construction ofequipment.

Generally, flameless combustion is accomplished by preheating combustionair and fuel gas (e.g., methane, methanol, hydrogen and the like)sufficiently such that when the two streams are combined the temperatureof the mixture exceeds the autoignition temperature of the mixture, butto a temperature less than that which would result in the oxidation uponmixing, being limited by the rate of mixing. Preheating of thecombustion air and fuel streams to a temperature between about 1500° F.and about 2300° F. and then mixing the streams in relatively smallincrements will result in flameless combustion. For some fuels such asmethanol, preheating to a temperature above about 1000° F. issufficient. The increments in which the fuel gas is mixed with thecombustion gas stream preferably result in about a 20° to about 200° F.temperature rise in the combustion gas stream due to the combustion ofthe fuel.

With most steam methane reforming processes controlling the temperaturein the catalyst bed is a problem. The advantages of the flamelessdistributed combustion as a heat source in the present process andapparatus can be summarized as follows:

-   -   FDC helps maintain a more uniform temperature, but        simultaneously controls heat-flux to match the local heat needed        for the material left to be reacted. At the highest heat flux        there is as much heat present as can be accommodated by the        reaction and as the process progresses less and less heat is        required to drive the reaction.    -   FDC has a lower maximum-temperature combustion gas.    -   FDC does not have hot spots which might damage the        hydrogen-selective, hydrogen-permeable membrane.    -   FDC has a negligible NO_(x) production.    -   FDC makes it easier to tailor axial heat flux distribution to        minimize entropy production or energy loss and, thus, making it        more efficient.    -   FDC permits a more compact reactor design that is less expensive        to build.    -   FDC permits a modular reactor design, at a wide range of sizes        and heat duties.

FDC provides a tapered heat flux profile.

Thus, the flameless distributed combustion (FDC) used to drive the steamreforming reactions in the present invention can be described ascomprising:

-   -   e) preheating either a fuel gas or oxidant or both to a        temperature that exceeds the autoignition temperature of the        mixture of the fuel gas and oxidant when they are mixed;    -   f) passing said fuel gas and oxidant in into a heating zone        which is in heat transferring contact along a substantial        portion of the reaction zone (i.e., the zone in which said        reforming reactions take place); and    -   g) mixing the fuel gas and oxidant in said heating zone in a        manner that autoignition occurs, resulting in combustion without        high temperature flames, thereby providing uniform, controllable        heat to said reaction zone.

In the practice of the invention, some degree of sulfur removal willprobably be necessary to protect the palladium material making up thehydrogen-permeable separation membrane and the Ni reforming catalyst.Sulfur is a temporary poison to such catalysts, but the catalystactivity can be regenerated by removing the source of sulfur. The sulfurtolerance of commercial reforming catalysts is dependent upon processconditions. On average, sulfur must be reduced to below 10 ppb to allowthe catalyst to function properly.

Feed clean up with ZnO beds or by other means known in the art may beused to remove impurities such as H₂S and other sulfur containingcompounds in the feed that could contribute to membrane degradation. Forheavier hydrocarbons, like naphtha, some hydrotreating may be necessaryto convert organic sulfur to H₂S, as known in the art. Heavy oil, solidscarried by liquid water, oxygen, amines, halides, and ammonia are alsoknown poisons for palladium membranes. Carbon monoxide competes withhydrogen for active surface sites, thus reducing the hydrogenpermeability by 10% at 3-5 Bar. Thus, the partial pressure needs to staylow for best performance, as is the case in our preferred design.

In another embodiment of the present invention the FDC-MSR generatedpure hydrogen is used in an integrated design to power a fuel cell. Thisembodiment of the present invention has the potential for about 71% orgreater efficiency in the generation of electricity from starting fuel.In addition, due to the unique integration of the system, CO₂ isproduced in high concentrations from about 80% to about 95% molar drybasis, and high pressure of from about 0.1 to about 20 MPa, particularlyfrom about 1 to about 5 MPa (S.I.), and is easier to separate fromnitrogen, which makes the system even more efficient.

Referring now to FIG. 6, a vaporizable hydrocarbon and steam 5 are fedinto the catalyst section 4 of a FDC-membrane reactor of the typedescribed in FIG. 1, while preheated air 7 and fuel 14 are fed into theFDC heating section 2 of the reactor containing fuel tubes 10. A sweepgas (in this case steam) is fed into the FDC-membrane reactor at 6. Theproduced high purity hydrogen stream, 12, is directed to the anodecompartment of the molten carbonate fuel cell, 20, operating at about650° C. and 5 Bar. The reactor effluent 13 containing the unreactedsteam, CO₂ and low quantities of methane, hydrogen and CO, and the fluegas 11 from the FDC heater and air, 16 are fed to the cathodecompartment of the same fuel cell, 17. The CO₂ reacts with the O₂ toform CO₃ ⁼ anions that transport through the molten carbonate membrane.The CO₃ ⁼ anions are constantly renewed. The reactions with indicatedtransport are described as follows:CO_(2 cathode)+½O_(2 cathode)+2e ⁻ _(cathode)→CO_(3 cathode) ⁼  R. 1CO_(3 cathode) ⁼→CO_(3 anode) ⁼  R. 2CO_(3 anode) ⁼→CO₂+½O_(2 anode)+2e ⁻ _(anode)  R. 3H_(2 anode)+½O_(2 anode)→H₂O_(anode)−242 kJ/gmol-H₂  R. 4Net: H_(2 anode)+½O_(2 cathode)+CO_(2 cathode+)2e ⁻_(cathode)→H₂O_(anode)+CO_(2 anode)+2e ⁻ _(anode)−242 kJ/gmol-H₂  R. 5

Electricity generated by the fuel cell is shown as electrical output 21.The stream from the anode, 22, now contains the permeated CO₂ and steambut no hydrogen, nitrogen, methane or oxygen, if hydrogen and oxygen arefed in exactly 2:1 stoichiometry. A portion of stream 22 may recycled tothe cathode compartment 17 of the fuel cell. The CO₂ recycle stream isshown as 23 on FIG. 6. A portion of streams 22 and/or 13 also may be putthrough a turbine expander to generate electrical or mechanical work 30and 24, respectively. In the present invention CO₂ is separated fromnitrogen essentially for free while electricity is simultaneouslygenerated. Furthermore the CO₂ capture leverage is high. As shown above,each mole of methane is converted to 4 moles of H₂. Therefore 4 moles ofCO₂ per mole of converted methane are required to transport the oxygenin the fuel cell and are therefore separated from the nitrogen. Thus,this process can also be used to separate CO₂ from an external CO₂containing stream. The high concentration CO₂ stream, 29, is now a primecandidate for sequestration after the steam is condensed. The CO₂ can beused for oil recovery, or injected into subterranean formations, orconverted to a thermodynamically stable solid. Also, since the presentprocess can be operated to produce high purity hydrogen and nitrogen aswell as concentrated CO₂, it can be used to facilitate the production ofchemicals such as urea, which can be made from these three rawmaterials. Other chemicals which can be manufactured using the productsand by-products of the present process include ammomia and ammoniumsulfate. Other uses for the concentrated stream of CO₂ and the highpurity hydrogen and nitrogen streams will be apparent to those skilledin the art.

The stream from the cathode, stream 18, contains all the nitrogen,unreacted oxygen, a little unpermeated CO₂, and trace amounts of themethane, hydrogen and CO from the MSR effluent. All or part of thisstream can be put through a turbine expander (not shown) to generatework (electrical or mechanical), 19. The trace components of stream 18may be oxidized in a catalytic converter, 26, and emitted in theatmosphere as a low CO₂ concentration containing stream, 27, containingless than 10% CO₂, preferably less than 1% CO₂. The trace components mayalso be oxidized inside the fuel cell if the appropriate catalyst isplaced in the cathode compartment. A stream, 28, containing water andsteam exits condenser 25 and is recycled to the FDC-MSR reactor, andreheated to between about 250 to 500° C.

The zero emission hybrid system of the present invention is extremelyefficient. Byproduct compounds are separated, the steam and hydrogen arereheated efficiently, and electricity is produced. Furthermore, water isseparated from purified CO₂ which is produced in concentrations largeenough to be easily sequestered. Advantages include using waste heat toraise steam and using water collected for recycling to supportadditional steam reforming or other beneficial uses. The system is atotally integrated, extremely efficient design having the potential forgreater than 71% generation efficiency as mentioned above. The 71% isapproximately a 20% fractional improvement over the best results we areaware of in the art, the 60% figure mentioned above that is possibleunder laboratory conditions. In addition to the great improvement inefficiency, the integrated design provides a concentrated source of CO₂for capture and sequestration as well.

Fuel cells which would be suitable for use in the present invention arethose that could function in a highly pressurized system. Most fuelcells run at atmospheric conditions. For this reason, a high pressuremolten carbonate fuel cell is preferred. However, other types of fuelcells, such as PEM fuel cells and SOFC, can also be effectively combinedwith the FDC-MSR reactor of the present invention.

Another very attractive feature is that the FDC powered MSR hydrogengenerator produces very low NO_(x), especially compared with thecombined processes known in the art. Due to the use of flamelessdistributed combustion very little NO_(x) is generated in this system.Furthermore, other steam reforming reactors used to generate hydrogenknown in the art could not feed to the MCFC the flue gas from thefurnace as in the present design, because they produce high NO_(x),which would poison the molten carbonate membrane.

The following illustrative embodiments will serve to illustrate theinvention disclosed herein. The examples are intended only as a means ofillustration and should not be construed as limiting the scope of theinvention in any way. Those skilled in the art will recognize manyvariations that may be made without departing from the spirit of thedisclosed invention.

ILLUSTRATIVE EMBODIMENT 1

FIG. 8 shows a schematic diagram of a multi-tubular, FDC heated, radialflow, membrane, steam reforming reactor in accordance with the presentinvention. In the reactor shown in FIG. 8, a vaporizable hydrocarbon andsteam enter the reactor at inlet 69 and flow through the reformingcatalyst bed 70 (which is in the form of an annulus) containing multiplemembrane tubes 71 and multiple FDC tubes 72 surrounded by the catalystbed. In this embodiment the feed gases and reaction gases flow throughthe catalyst bed radially from outside to inside. The multiplehydrogen-selective, hydrogen-permeable, membrane tubes 71 are disposedaxially in concentric rows in the reforming catalyst bed and serve toremove hydrogen, which is produced by the reforming reactions. Themultiple FDC tubes (i.e., chambers) 72 are also disposed axially inconcentric rows in the reforming catalyst bed (for example, in a ratioof 1:2 or other number of FDC tubes to the number of membrane tubes).The multiple FDC tubes are in contact with the reforming catalyst bedand provide a controlled, distributed heat flux to the catalyst bedsufficient to drive the reforming reactions. While the membrane tubesand the FDC tubes are shown to be in concentric rows in FIG. 8, othergeometric arrangements of these tubes can be suitably employed, and arewithin the scope of the present invention.

The FDC tubes 72 generally comprise a fuel conduit disposed within alarger tube with an inlet and flow path for a preheated oxidant (e.g.,preheated air) and an outlet for combustion (flue) gas. The FDC tubesmay be closed ended with a fuel conduit, oxidant inlet and flow path,and flue gas outlet arranged as shown in FIG. 10A, or may open endedwith the fuel conduit, oxidant inlet and flow path arranged as shown inFIG. 10B.

High purity hydrogen is removed from the multi-tubular, radial flow,reactor shown in FIG. 8 via outlets 73, with the aid of vacuum.Optionally, a sweep gas may be used to promote the diffusion of hydrogenthrough the membrane of the membrane tubes 71. If a sweep gas isemployed, the membrane tubes 71 may contain an outer sweep gas feed tubeand an inner return tube for sweep gas and hydrogen as discussed in FIG.12. By-product gases, including uppermeated hydrogen, if not furtherused internally for heat production, e.g., combustion or heat exchange,exit the multi-tubular, radial flow, reactor via outlet 74. A hollowtube or cylinder 75 may optionally be used for flow distribution.

ILLUSTRATIVE EMBODIMENT 2

FIG. 9 is a top cross-section view of the shell of the multi-tubular,FDC heated, radial flow, membrane, steam reforming reactor of FIG. 8.The cross sectional view of the reactor shows multiple membrane tubes 71and multiple FDC tubes 72 dispersed in catalyst bed 70 with optionalhollow tube or cylinder 75 being in the center of the reactor. In theexample shown, the membrane tubes 71 have outside diameters (OD) ofabout one inch while FDC tubes have an OD of approximately two inches,although other sizes of these tubes can be suitably employed. If a sweepgas is employed, the membrane tubes 71 may contain an outer sweep gasfeed tube and an inner return tube for sweep gas and hydrogen as shownin FIGS. 12 and 14. A larger shell containing more tubes duplicatingthis pattern can also be used.

ILLUSTRATIVE EMBODIMENT 3

FIGS. 10A and 10B are schematic diagrams showing an example of a “closedended” and of an “open ended” FDC tubular chamber which are used todrive the reforming reactions in various embodiments of the presentinvention. Referring to FIG. 10A, an oxidant (in this case preheatedair) enters the FDC tube at inlet 76 and mixes with fuel which entersthe FDC tube at inlet 77 and passes into fuel conduit 78 through nozzles79 spaced along the length of the fuel conduit, whereupon it mixes withthe air which has been preheated to a temperature such that thetemperature of the resulting mixture of fuel and air is above theautoignition temperature of the mixture. The reaction of the fuelpassing through the nozzles and mixing with the flowing preheated air ata temperature above the autoignition temperature of the mixture, resultsin flameless distributed combustion which releases controlled heat alongthe length of the FDC tube as shown, with no flames or hot spots. Thecombustion gases, (i.e., flue gas) exit the FDC tube at outlet 80.

In the “open ended” FDC tubular chamber shown in FIG. 10B, preheated airenters the FDC tube at inlet 76 and the fuel at inlet 77, and the fuelpasses through conduit 78 and nozzles 79, similar to “closed end” FDCtube in FIG. 10A. However, in the case of the “open ended” FDC tube, theflue gas exits the FDC tube at open end 81, instead of outlet 80 asshown in FIG. 10A.

ILLUSTRATIVE EMBODIMENT 4

FIG. 11 is a schematic drawing of a multi-tubular, FDC heated, axialflow, membrane, steam reforming reactor in accordance with the presentinvention. In the reactor shown in FIG. 11, a vaporizable hydrocarbonand steam enter the reactor at inlet 69 and flow through the reformingcatalyst bed 70 containing multiple hydrogen-selective membrane tubes 71and multiple FDC tubes 72. In this embodiment the feed gases andreaction gases flow through the catalyst bed axially from the top of thecatalyst bed to the bottom. The multiple hydrogen-selective membranetubes 71 are disposed axially in the reforming catalyst bed and serve toremove hydrogen which is produced by the reforming reactions. In theembodiment shown the membrane tubes are closed at the top and a sweepgas (e.g. steam) is employed, which enters the reactor at inlet 85 intothe bottom of the membrane tubes where it flows upward in the outer partof the membrane tube, counter-current to the hydrocarbon and steam feed.A stinger pipe fitted to the bottom of the permeate section may be usedto distribute the sweep gas in the membrane tube. The permeated hydrogenand sweep gas flow downward in a return tube located in the center ofthe membrane tube and exit the reactor via outlet 86. The pressure dropin the permeate pipe section is significant when the length of the piperelative to the diameter exceeds a given limit. Actually, the volumetricamount of hydrogen crossing the membrane is proportional to the membranearea, □*D*L and the multiplier is the velocity, which is fixed as afunction relating to Sievert's law, the description of which can befound in US2003/0068269 and is herein incorporated by reference. Thesame hydrogen amount has to flow across the pipe cross section which isequal to □*D²/4. The ratio of hydrogen velocities through the pipe andthrough the membrane respectively is proportional to (□*D*L)/ . . .□*D²/4) or to L/D. Pressure drop increases with gas velocity. If thisratio exceeds a limit, then the velocity in the permeate pipe exceeds alimit too, since the velocity through the membrane is fixed. Then thepressure drop in the permeate pipe becomes high and it reduces thehydrogen flux by creating back pressure in the permeate section. In sucha case, the reactor design has to accommodate either a higher membranediameter, or a reduced length.

There are also multiple FDC tubes (i.e., chambers) 72 disposed axiallyin the reforming catalyst bed. In the embodiment shown the FDC tubes are“closed ended” tubes with preheated air entering at inlet 76, fuelentering at 77 and combustion gas (i.e., flue gas) exiting the reactorat outlet 80. The multiple FDC tubes are in heat transferring contactwith the reforming catalyst bed 70 and provide a controlled, distributedheat flux to the catalyst bed sufficient to drive the reformingreactions. While the membrane tubes and the FDC tubes are shown to be ina particular geometric pattern in FIG. 11, it is understood that othergeometric arrangements of these tubes may be used and are within thescope of the invention. While “closed ended” FDC tubes are employed inthe particular reactor shown in FIG. 11, “open ended” FDC tubes may besuitably employed as well. Also, the FDC tubes and/or the membrane tubesmay be surrounded by cylindrical screens (not shown) to protect themfrom getting in direct contact with the catalyst, and allow insertion ofthese tubes even after the catalyst is loaded into the reactor.

The FDC chaser must be free of obstructions and have a tubular dimensionfor the external or exterior tube of the FDC chamber such that thelength to diameter ratio is higher than a given limit, preferably morethan 4. This ratio ensures that the air velocity in the chamber becomeshigher than the flame velocity of the fuel and that turbulence isinduced to improve heat transfer. In such a condition, no flames arecreated or stabilized. Any obstructions (like baffles) would createstagnation points where flames would form and stabilize.

High purity hydrogen, which diffuses through the membrane into themembrane tubes, is removed from the reactor via outlet(s) 86 togetherwith the sweep gas (in this case steam). While outlet 86 is shown inFIG. 11 to be located on the side of the reactor, this outlet mayoptionally be located at the bottom of the reactor thereby avoiding abottom side exit manifold. A further option involves the use of a vacuuminstead of a sweep gas to facilitate diffusion of the hydrogen throughthe membrane into the membrane tubes. Vacuum can be induced eithermechanically with a pump or chemically with a metal hydride precursorwhich reacts away the hydrogen to form metal hydride. The hydride ison-line for a given period of time and when it is saturated, a parallelcompartment can be put on-line, while the original compartment isisolated and heated to desorb and produce the hydrogen. This isadvantageous in cases where the hydrogen needs to be stored and/orshipped to a customer or in cases where the cost of electrical energyfor running a pump is higher than using waste energy to desorb thehydrogen from the hydride. Detailed economics will dictate the rightchoice.

In another embodiment of the reactor in FIG. 11, the sweep gas inlet 85and the hydrogen, sweep gas outlet 86 and their associated plenums, maybe placed on the top of the reactor allowing easy access to the bottomof the reactor. In a further embodiment of the reactor of FIG. 11, thepreheated air inlet 76, the fuel inlet 77 and the flue gas outlet 80 andtheir associated plenums may be placed on the bottom of the reactorallowing easy access to the top of the reactor.

By-product gases, including carbon dioxide, steam, and minor amounts ofcarbon monoxide and unpermeated hydrogen, if not further used internallyfor heat production, e.g., combustion or heat exchange, exit themulti-tubular, axial flow, reactor via outlet 74. The reactor shown inFIG. 11 may be equipped with baffles and/or screens such as the bafflesshown in FIGS. 13A and 13B or 13C and 13D.

ILLUSTRATIVE EMBODIMENT 5

FIG. 12 is a top cross-section view of the shell of the multi-tubular,FDC heated, axial flow, membrane reactor shown in FIG. 11. In theembodiment shown multiple membrane tubes 71 and multiple FDC tubes 72are dispersed in reforming catalyst bed 70. The multiple FDC tubesemployed in this embodiment are “closed ended” FDC tubes as discussedabove in connection with FIG. 11. The membrane tubes are equipped withan outer sweep gas feed tube and an inner hydrogen, sweep gas returntube as discussed in connection with FIG. 11. A typical reactor of thetype shown in this FIG. 12 may comprise, for example, 19 FDC tubes of5.5″ OD and 90 membrane tubes of 2″ OD enclosed in a shell of 3.5 ftdiameter containing catalyst in the void spaces. Other shell sizes andnumbers of tubes can be suitably employed depending on the capacityneeded. The design parameter which is of outmost importance is theoptimum gap between the membrane and the FDC tubes. If a high gap isassumed, then heat transfer limitations occur since the flow of enthalpyfrom FDC to the reforming reaction is slow. The membranes may notoperate isothermally and cold spots may develop, thus reducing thereactor efficiency. If a small gap is assumed, then there may beproblems with insufficient catalyst penetration in the gap, overheatingof the membrane, or even touching of the hot FDC tube with the membranein conditions where the tubes are not perfectly straight. A narrow gaplimitation will make reactor fabrication more expensive, sinceclearances are hard to achieve. Thus, an intermediate gap is morepreferable. As a particular non-limiting example, the gap between themembrane and the FDC tubes is from about ¼ inch (about 0.64 cm) to about2 inches (about 5.08 cm), particularly from about ½ inch (about 1.27 cm)to about 1 inch (about 2.54 cm). The gap between the membrane tubes maybe from about ¼ inch to about 2 inches, particularly from about ½ inchto about 1 inch and this has to be also optimized. Thehydrogen-permeable membrane tube has a ratio of length to diameter ofless than about 500.

ILLUSTRATIVE EMBODIMENT 6

FIGS. 13A and 13B and 13C and 13D show two different configurations ofbaffles which may be employed in the multi-tubular, FDC heated, axialflow, membrane steam reforming reactors of the invention to increasecontact of the reactant gases with the catalyst in the catalyst beds.The baffle configuration shown in FIGS. 13A and 13B comprise a washershaped baffle 87 and a disk shaped baffle 88 arranged in and alternatingpattern. This baffle arrangement causes the feed and reactant gases toflow through the hole in the washer shaped baffle and be deflected bydisk shaped baffle thereby enhancing the contact of the reactant gaseswith the catalyst (not shown) which is packed in the area between thebaffles.

The baffle arrangement shown in FIGS. 13C and 13D comprises truncateddisks 89 which are placed in and alternating pattern (truncated left andtruncated right) in the reactor thereby causing the feed and reactantgases to “zigzag” as they flow through the catalyst (not shown) which ispacked in the area between the baffles.

The baffles in FIGS. 13A&B and 13C&D will have openings (not shown) toallow the FDC tubes and membrane tubes to pass through them. Screenspositioned in vertical alignment (not shown) may also be used to supportthe baffles and in some cases hold the catalyst away from the shell wallor from the center of the shell for better gas flow distribution.

ILLUSTRATIVE EMBODIMENT 7

FIG. 14 is a top cross-section view of the shell of a multi-tubularreactor in accordance with one embodiment of the invention in which fourmembrane tubes 71 are dispersed in the reforming catalyst bed 70 whichis packed into reactor tube 82, while the FDC chamber is in the form ofan annulus surrounding the reforming catalyst bed. The tubular FDCchamber (which is defined by outer wall 83 and the wall of the reactortube 82) contains multiple fuel conduits 78 having nozzles (not shown)through which fuel flows and mixes with preheated air flowing in the FDCchamber whereupon flameless combustion occurs. If a sweep gas isemployed, the membrane tubes 71 may contain an outer sweep gas feed tubeand an inner return tube for sweep gas and hydrogen as shown in FIG. 14.In one embodiment of the invention, the membrane tubes have an OD of 2inches, while the outer FDC tube has an inner diameter (ID) ofapproximately 8.6 inches. However, other sizes can be suitably employed.

ILLUSTRATIVE EMBODIMENT 8

FIG. 15 is a top cross-section view of the shell of another embodimentof the multi-tubular, axial flow, reactor of the invention in whichmultiple reactor tubes 82 packed with reforming catalyst are employed.In this example each of the six reactor tubes 82 contains a catalyst bed70 and a membrane tube 71 containing an outer sweep gas feed tube and aninner hydrogen, sweep gas return tube. Heat is provided to the reformingcatalyst beds by the tubular FDC chamber defined by outer wall 83 andinner wall 84. The FDC chamber contains multiple fuel conduits 78dispersed at various intervals in the FDC chamber. A hollow tube orcylinder defined by inner wall 84 may optionally be used for flowdistribution.

ILLUSTRATIVE EMBODIMENT 9

FIG. 16 is a top cross-section view of the shell of further embodimentof the multi-tubular, axial flow, reactor of the invention in which fourmembrane tubes are dispersed in each of six reactor tubes 82 containingcatalyst beds 70. Heat is provided to the catalyst beds by FDC chamberdefined by outer wall 83 and inner wall 84. The FDC chamber containsmultiple fuel conduits 78 having nozzles 79 (not shown). If a sweep gasis employed, the membrane tubes 71 may contain an outer sweep gas feedtube and an inner return tube for sweep gas and hydrogen as discussedshown and discussed above in connection with FIGS. 12 and 14. The hollowcylinder or tube defined by inner wall 84 may optionally be used forflow distribution.

ILLUSTRATIVE EMBODIMENT 10

FIG. 17 is a top cross-section view of the shell of further embodimentof the multi-tubular, axial flow, reactor of the invention in which sixmembrane tubes 71 are dispersed in each of the six reactor tubes 82packed with reforming catalyst. Heat is provided to the reformingcatalyst beds by the FDC chamber defined by outer wall 83 and inner wall84. The FDC chamber contains multiple fuel conduits 78. Additional heatmay be provided to the catalyst beds by employing an FDC tube 72 in thecenter of each of the reactor tubes 82 as shown in FIG. 17. The hollowtube or cylinder defined by inner wall 84 may optionally be used forflow distribution.

If a sweep gas is employed, the membrane tubes 71 may contain an outersweep gas feed tube and an inner return tube for sweep gas and hydrogenas discussed in FIG. 12.

Other illustrative embodiments include Examples 1-6 of US2003/0068269and the description thereof is herein incorporated by reference.

The ranges and limitations provided in the instant specification andclaims are those, which are believed to particularly point out anddistinctly claim the instant invention. It is, however, understood thatother ranges and limitations that perform substantially the samefunction in substantially the same manner to obtain the same orsubstantially the same result are intended to be within the scope of theinstant inventions defined by the instant specification and claims

1. A flameless distributed combustion heated; membrane, steam reformingreactor comprising: a) a reforming chamber containing a reformingcatalyst bed, said reforming chamber having an inlet for vaporizablehydrocarbon and steam, a flow path for hydrogen and by-product gasesresulting from the reforming reactions taking place in said reformingchamber and an outlet for said by-product gases, b) at least oneflameless distributed combustion chamber in a heat transferringrelationship with said reforming catalyst bed whereby a distributed,controlled heat flux is provided by said flameless distributedcombustion chambers(s) to said reforming catalyst bed, said flamelessdistributed combustion chamber(s) comprising an inlet and a flow pathfor an oxidant, an outlet for combustion gas and further comprising afuel conduit having an inlet for fuel and a plurality of fuel nozzleswhich provide fluid communication from within the fuel conduit to theflow path of said oxidant, said plurality of fuel nozzles being sizedand spaced along the length of said fuel conduit so that no flameresults when said fuel is mixed with said oxidant in said flamelessdistributed combustion chamber; c) a preheater capable of preheatingsaid oxidant to a temperature that when said fuel and said oxidant aremixed in said flameless distributed combustion chamber, the temperatureof the resulting mixture of said oxidant and fuel exceeds theautoignition temperature of said mixture; and d) at least twohydrogen-selective, hydrogen-permeable, membrane tubes in contact withsaid reforming catalyst bed, each of said membrane tubes having anoutlet whereby hydrogen formed in the reforming chamber permeates intosaid membrane tube and passes through said outlet.
 2. A process for theproduction of hydrogen, comprising: a) reacting steam with a vaporizablehydrocarbon at a temperature of from about 200° C. to about 700° C. andat a pressure of from about 1 bar to about 200 bar in a reaction zonecontaining a reforming catalyst to produce a mixture of primarilyhydrogen and carbon dioxide, with a lesser amount of carbon monoxide; b)providing heat to said reaction zone by employing at least one flamelessdistributed combustion chamber thereby driving said reaction; and c)conducting said reaction in the vicinity of at least twohydrogen-permeable, hydrogen-selective membrane tubes, whereby hydrogenformed in said reaction zone permeates through said hydrogen selectivemembrane tubes and is separated from said carbon dioxide and carbonmonoxide.
 3. A membrane, steam reforming reactor comprising: a) areforming chamber containing a reforming catalyst bed, said reformingchamber having an inlet for vaporizable hydrocarbon and steam, a flowpath for hydrogen and by-product gases resulting from the reformingreactions taking place in said reforming chamber and an outlet for saidby-product gases, b) at least one flameless distributed combustionchamber in a heat transferring relationship with said reforming catalystbed, and c) at least two hydrogen-selective, hydrogen-permeable,membrane tubes in contact with said reforming catalyst bed, wherein atleast one of the membrane tubes is connected to a section containing ametal hydride precursor, and wherein the hydrogen formed in thereforming chamber permeates through said membrane tube to said sectioncontaining the metal hydride precursor which reacts with the permeatedhydrogen to form hydride.
 4. (canceled)
 5. (canceled)
 6. The hydrogenfuel cell, wherein the hydrogen feed is made by a process as describedin claim
 2. 7. (canceled)
 8. The process of claim 2, wherein saidcatalyst bed is in heat transferring contact with multiple flamelessdistributed combustion chambers.
 9. (canceled)
 10. (canceled)
 11. Theprocess of claim 2, wherein a sweep gas is used to promote the diffusionof hydrogen through at least one of said membrane tubes, said sweep gasbeing selected from the group consisting of steam, carbon dioxide,nitrogen and condensable hydrocarbon and the vaporizable hydrocarbon isselected from the group consisting of natural gas, methane, ethylbenzene, methanol, ethane, ethanol, propane, butane, light hydrocarbonshaving 1-4 carbon atoms in each molecule, light petroleum fractionsincluding naphtha, diesel, kerosene, jet fuel or gas oil, and hydrogen,carbon monoxide and mixtures thereof.
 12. (canceled)
 13. The reactor, ofclaims 1, wherein said catalyst bed contains baffles in a form selectedfrom the group consisting of (i) washers and disks, and (ii) truncateddisks.
 14. The reactor, of claim 1, wherein the hydrogen-selective andat least one of the hydrogen-permeable membranes is made of a Pd-alloylayer supported on a porous metal, particularly a Pd-alloy layerdeposited by electroless plating on porous metal with an in-situ oxideprotection layer.
 15. (canceled)
 16. The reactor, of claim 1, wherein atleast one of the hydrogen-selective and hydrogen-permeable membranes hasa ratio of length to diameter of less than about 500, gaps between themembrane tubes are from about ¼ inch (about 0.64 cm) to about 2 inches(about 5.08 cm), and gap between the membrane and FDC tubes is fromabout ¼ inch (about 0.64 cm) to about 2 inches (about 5.08 cm). 17.(canceled)
 18. The reactor, of claim 1, wherein the FDC chamber has anexternal tubular dimension such that the length to diameter ratio ishigher than
 4. 19. (canceled)
 20. (canceled)
 21. The process of claim 2,wherein carbon dioxide produced from said steam reforming chamber has aconcentration of from about 80% to about 99% molar dry basis. 22.(canceled)
 23. The process of claim 2, wherein carbon dioxide producedfrom said steam reforming chamber is used at least in part for enhancedrecovery of oil in oil wells or enhanced recovery of methane in coal bedmethane formations.
 24. The reactor of claim 1, wherein said catalystbed is in heat transferring contact with multiple flameless distributedcombustion chambers.
 25. The reactor of claim 3, wherein said catalystbed is in heat transferring contact with multiple flameless distributedcombustion chambers.
 26. The reactor of claim 3, wherein a sweep gas isused to promote the diffusion of hydrogen through at least one of saidmembrane tubes, said sweep gas being selected from the group consistingof steam, carbon dioxide, nitrogen and condensable hydrocarbon and thevaporizable hydrocarbon is selected from the group consisting of naturalgas, methane, ethyl benzene, methanol, ethane, ethanol, propane, butane,light hydrocarbons having 1-4 carbon atoms in each molecule, lightpetroleum fractions including naphtha, diesel, kerosene, jet fuel or gasoil, and hydrogen, carbon monoxide and mixtures thereof.
 27. The reactorof claim 3, wherein said catalyst bed contains baffles in a formselected from the group consisting of (i) washers and disks, and (ii)truncated disks.
 28. The reactor of claim 3, wherein thehydrogen-selective and at least one of the hydrogen-permeable membranesis made of a Pd-alloy layer supported on a porous metal, particularly aPd-alloy layer deposited by electroless plating on porous metal with anin-situ oxide protection layer.
 29. The reactor of claim 3, wherein atleast one of the hydrogen-selective and hydrogen-permeable membranes hasa ratio of length to diameter of less than about 500, gaps between themembrane tubes are from about 14 inch (about 0.64 cm) to about 2 inches(about 5.08 cm), and gap between the membrane and FDC tubes is fromabout 14 inch (about 0.64 cm) to about 2 inches (about 5.08 cm).
 30. Thereactor of claim 3, wherein the FDC chamber has an external tubulardimension such that the length to diameter ratio is higher than 4.